Split feed addition to iso-pressure open refrigeration lpg recovery

ABSTRACT

A process is disclosed herein for recovery of natural gas liquids from a feed gas stream, comprising forming a first portion of the feed gas stream and a second portion of the feed gas stream, wherein the mass ratio of the first portion to the second portion is in the range of 95:5 to 5:95, cooling the first portion in a heat exchanger and at least partially condensing the first portion, and feeding the second portion and the cooled and at least partially condensed first portion to a distillation column wherein lighter components are removed from the distillation column as an overhead vapor stream and heavier components are removed from the distillation column in the bottoms as a product stream, and wherein the second portion is fed into the distillation column at a point one or more vapor-liquid equilibrium stages below the first portion, thereby allowing mass transfer exchange between liquids of the cooled second portion and vapors of the second portion within the column. A corresponding apparatus is also disclosed.

RELATED APPLICATIONS

This application claims priority from U.S. Provisional Application No. 61/888,901 filed Oct. 9, 2013.

FIELD

The embodiments described herein relate to improved processes for recovery of natural gas liquids from gas feed streams containing hydrocarbons, and in particular to recovery of propane and ethane from gas feed streams.

BACKGROUND

Natural gas contains various hydrocarbons, including methane, ethane and propane.

Natural gas usually has a major proportion of methane and ethane, i.e. methane and ethane together typically comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, carbon dioxide and other gases. In addition to natural gas, other gas streams containing hydrocarbons may contain a mixture of lighter and heavier hydrocarbons. For example, gas streams formed in the refining process can contain mixtures of hydrocarbons to be separated. Separation and recovery of these hydrocarbons can provide valuable products that may be used directly or as feedstocks for other processes. These hydrocarbons are typically recovered as natural gas liquids (NGL).

The embodiments described herein are primarily directed to recovery of C₃+ components in gas streams containing hydrocarbons, and in particular to recovery of propane from these gas streams. A typical natural gas feed to be processed in accordance with the processes described below typically may contain, in approximate mole percent, 92.12% methane, 3.96% ethane and other C₂ components, 1.05% propane and other C₃ components, 0.15% iso-butane, 0.21% normal butane, 0.11% pentanes or heavier, and the balance made up primarily of nitrogen and carbon dioxide. Refinery gas streams may contain less methane and higher amounts of heavier hydrocarbons.

Recovery of natural gas liquids from a gas feed stream has been performed using various processes, such as cooling and refrigeration of gas, oil absorption, refrigerated oil absorption or through the use of multiple distillation towers. More recently, cryogenic expansion processes utilizing Joule-Thompson valves or turbo expanders have become preferred processes for recovery of NGL from natural gas.

In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high pressure liquids containing the desired components.

The high-pressure liquids may be expanded to a lower pressure and fractionated. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation column. In the distillation column volatile gases and lighter hydrocarbons are removed as overhead vapors and heavier hydrocarbon components exit as liquid product in the bottoms.

The feed gas is typically not totally condensed, and the vapor remaining from the partial condensation may be passed through a Joule-Thompson valve or a turbo expander to a lower pressure at which further liquids are condensed as a result of further cooling of the stream. The expanded stream is supplied as a feed stream to the distillation column.

A reflux stream is provided to the distillation column, typically a portion of partially condensed feed gas after cooling but prior to expansion. Various processes have used other sources for the reflux, such as a recycled stream of residue gas supplied under pressure.

While various improvements to the general cryogenic processes described above have been attempted, these improvements continue to use a turbo expander or Joule-Thompson valve to expand the feed stream to the distillation column. It would be desirable to have an improved process for enhanced recovery of NGLs from a natural gas feed stream.

SUMMARY

The embodiments described herein relate to improved processes for recovery of NGLs from a feed gas stream. The process utilizes an open loop mixed refrigerant process to achieve the low temperatures necessary for high levels of NGL recovery. A single distillation column is utilized to separate heavier hydrocarbons from lighter components such as sales gas. The overhead stream from the distillation column is cooled to partially liquefy the overhead stream. The partially liquefied overhead stream is separated into a vapor stream comprising lighter hydrocarbons, such as sales gas, and a liquid component that serves as a mixed refrigerant. The mixed refrigerant provides process cooling and a portion of the mixed refrigerant is used as a reflux stream to enrich the distillation column with key components. With the gas in the distillation column enriched, the overhead stream of the distillation column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than typically used for high recoveries of NGLs. The process achieves high recovery of desired NGL components without expanding the gas as in a Joule-Thompson valve or turbo expander based plant, and with only a single distillation column.

In one embodiment, C₃+ hydrocarbons, and in particular propane, are recovered. Temperatures and pressures are maintained as required to achieve the desired recovery of C₃+ hydrocarbons based upon the composition of the incoming feed stream. In this embodiment of the process, feed gas enters a main heat exchanger and is cooled. The cooled feed gas is fed to a distillation column, which in this embodiment functions as a deethanizer. Cooling for the feed stream may be provided primarily by a warm refrigerant such as propane. The overhead stream from the distillation column enters the main heat exchanger and is cooled to the temperature required to produce the mixed refrigerant and to provide the desired NGL recovery from the system.

The cooled overhead stream from the distillation column is combined with an overhead stream from a reflux drum and separated in a distillation column overhead drum. The overhead vapor from the distillation column overhead drum is sales gas (i.e. methane, ethane and inert gases) and the liquid bottoms are the mixed refrigerant. The mixed refrigerant is enriched in C₂ and lighter components as compared to the feed gas. The sales gas is fed through the main heat exchanger where it is warmed. The temperature of the mixed refrigerant is reduced to a temperature cold enough to facilitate the necessary heat transfer in the main heat exchanger. The temperature of the refrigerant is lowered by reducing the refrigerant pressure across a control valve. The mixed refrigerant is fed to the main heat exchanger where it is evaporated and super-heated as it passes through the main heat exchanger.

After passing through the main heat exchanger, the mixed refrigerant is compressed. Preferably, the compressor discharge pressure is greater than the distillation column pressure so no reflux pump is necessary. The compressed gas passes through the main heat exchanger, where it is partially condensed. The partially condensed mixed refrigerant is routed to a reflux drum. The bottom liquid from the reflux drum is used as a reflux stream for the distillation column. The vapors from the reflux drum are combined with the distillation column over head stream exiting the main heat exchanger and the combined stream is routed to the distillation column overhead drum. In this embodiment, the process can achieve over 99 percent recovery of propane from the feed gas.

In another embodiment of the process, the feed gas is treated as described above and a portion of the mixed refrigerant is removed from the plant following compression and cooling. The portion of the mixed refrigerant removed from the plant is fed to a C₂ recovery unit to recover the ethane in the mixed refrigerant. Removal of a portion of the mixed refrigerant stream after it has passed through the main heat exchanger and been compressed and cooled has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used as a feed stream in an ethylene cracking unit.

In another embodiment of the process, an absorber column is used to separate the distillation column overhead stream. The overhead stream from the absorber is sales gas, and the bottoms are the mixed refrigerant.

In yet another embodiment, only one separator drum is used. In this embodiment, the compressed, cooled mixed refrigerant is returned to the distillation column as a reflux stream.

The process described above may be modified to achieve separation of hydrocarbons in any manner desired. For example, the plant may be operated such that the distillation column separates C.sub.4+ hydrocarbons, primarily butane, from C₃ and lighter hydrocarbons. In another embodiment, the plant may be operated to recover both ethane and propane. In this embodiment, the distillation column is used as a demethanizer, and the plant pressures and temperatures are adjusted accordingly. In this embodiment, the bottoms from the distillation tower contain primarily the C₂+ components, while the overhead stream contains primarily methane and inert gases. In this embodiment, recovery of as much as 55 percent of the C₂+ components in the feed gas can be obtained.

Among the advantages of the process is that the reflux to the distillation column is enriched, for example in ethane, reducing loss of propane from the distillation column. The reflux also increases the mole fraction of lighter hydrocarbons, such as ethane, in the distillation column making it easier to condense the overhead stream. This process uses the liquid condensed in the distillation column overhead twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column. Other advantages of the processes described herein will be apparent to those skilled in the art based upon the detailed description of preferred embodiments provided below.

In yet another embodiment, a process is provided for recovery of natural gas liquids from a feed gas stream, comprising forming a first portion of the feed gas stream and a second portion of the feed gas stream, wherein the mass ratio of the first portion to the second portion is in the range of 95:5 to 5:95; cooling the first portion in a heat exchanger and at least partially condensing the first portion; and feeding the second portion and the cooled and at least partially condensed first portion to a distillation column wherein lighter components are removed from the distillation column as an overhead vapor stream and heavier components are removed from the distillation column in the bottoms as a product stream, and wherein the second portion is fed into the distillation column at a point one or more vapor-liquid equilibrium stages below the first portion, thereby allowing mass transfer exchange between liquids of the cooled first portion and vapors of the second portion within the column. The process further includes feeding the distillation column overhead stream to the heat exchanger and cooling the distillation column overhead stream to at least partially liquefy the distillation column overhead stream, feeding the at least partially liquefied distillation column overhead stream to a first separator, separating the vapor and liquid in the first separator to produce an overhead vapor stream comprising sales gas and a bottoms stream comprising a mixed refrigerant, feeding the mixed refrigerant stream to the heat exchanger to provide cooling, wherein the mixed refrigerant stream vaporizes as it passes through the heat exchanger, compressing the vaporized mixed refrigerant stream and passing the compressed mixed refrigerant stream through the heat exchanger, and feeding at least a portion of the compressed mixed refrigerant stream to the distillation column as a reflux stream. In embodiments, energy inputs are about 5-30% lower, or about 10-20% lower, than the energy inputs for processes in which the feed stream is not split and the entire feed stream passes through the heat exchanger for cooling. The decrease in energy input results in significant savings in operational costs.

In a further embodiment, an apparatus is provided for separating natural gas liquids from a feed gas stream, the apparatus comprising a primary feed gas line configured to deliver a feed gas stream, a heat exchanger operable to provide the heating and cooling necessary for separation of natural gas liquids from a feed gas stream by heat exchange contact between the feed gas stream and one or more process streams thus forming a cooled feed gas stream, and a distillation column configured to receive the feed gas stream and to separate the feed gas stream into a column overhead stream comprising a substantial amount of the lighter hydrocarbon components of the feed gas stream and a column bottoms stream comprising a substantial amount of the heavier hydrocarbon components. The apparatus further includes a first separator configured to receive the distillation column overhead stream and to separate the column overhead stream into an overhead sales gas stream and a bottoms stream comprising a mixed refrigerant configured to provide process cooling in the heat exchanger, a compressor configured to compress the mixed refrigerant stream after the mixed refrigerant stream has provided process cooling in the heat exchanger, and a feed gas bypass line configured to remove a portion of the feed gas stream prior to it being sent to the heat exchanger, wherein the feed gas bypass line is fluidly connected to the distillation column at a point one or more vapor-liquid equilibrium stages below the point at which the cooled feed gas stream from the heat exchanger is fluidly connected, thereby allowing mass transfer exchange between the liquids of the cooled feed gas stream from the heat exchanger and the vapors of the bypass feed gas stream within the column.

DESCRIPTION OF THE FIGURES

FIG. 1 is a schematic drawing of a plant for performing embodiments of a method in which the mixed refrigerant stream is compressed and returned to the reflux separator.

FIG. 2 is a schematic drawing of a plant for performing embodiments of a method in which a portion of the compressed mixed refrigerant stream is removed from the plant for ethane recovery.

FIG. 3 is a schematic drawing of a plant for performing embodiments in which an absorber is used to separate the distillation overhead stream.

FIG. 4 is a schematic drawing of a plant for performing embodiments in which only one separator drum is used.

FIG. 5 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and the mixed refrigerant is compressed and returned to the reflux separator.

FIG. 6 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and a portion of the compressed mixed refrigerant stream is removed from the plant for ethane recovery.

FIG. 7 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and an absorber is used to separate the distillation column overhead stream.

FIG. 8 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and in which only one separator drum is used.

FIG. 9 is a schematic drawing of a plant for performing embodiments of another method in which the feed stream to the distillation column is split and fed to different locations of the column.

DETAILED DESCRIPTION OF EMBODIMENTS

The embodiments described herein relate to improved processes for recovery of natural gas liquids (NGL) from gas feed streams containing hydrocarbons, such as natural gas or gas streams from petroleum processing. In embodiments, the process runs at approximately constant pressures with no intentional reduction in gas pressures through the plant. The process uses a single distillation column to separate lighter hydrocarbons and heavier hydrocarbons. An open loop mixed refrigerant provides process cooling to achieve the temperatures required for high recovery of NGL gases. The mixed refrigerant is comprised of a mixture of the lighter and heavier hydrocarbons in the feed gas, and is generally enriched in the lighter hydrocarbons as compared to the feed gas.

The open loop mixed refrigerant is also used to provide an enriched reflux stream to the distillation column, which allows the distillation column to operate at higher temperatures and enhances the recovery of NGLs. The overhead stream from the distillation column is cooled to partially liquefy the overhead stream. The partially liquefied overhead stream is separated into a vapor stream comprising lighter hydrocarbons, such as sales gas, and a liquid component that serves as a mixed refrigerant.

In embodiments, the process may be used to obtain the desired separation of hydrocarbons in a mixed feed gas stream. In one embodiment, the process of the present application may be used to obtain high levels of propane recovery. Recovery of as much as 99 percent or more of the propane in the feed case may be recovered in the process. The process can also be operated in a manner to recover significant amounts of ethane with the propane or reject most of the ethane with the sales gas. Alternatively, the process can be operated to recover a high percentage of C.sub.4+ components of the feed stream and discharge C₃ and lighter components.

In embodiments, a substantial reduction in energy usage can be obtained using the split feed configuration described herein. In embodiments, compressor duty can be reduced by at least 5%, or at least 10, or by 5-20% as compared to a system in which a split feed is not used. In embodiments, the reboiler duty can be reduced by at least 10%, or at least 20%, or at least 30% as compared to a system that does not have a split feed. In embodiments, the size of the distillation column also can be reduced, resulting in lower capital cost. In embodiments, the mass ratio of the first portion of the feed stream, which is cooled and partially condensed before being fed to the distillation column, and the second portion of the feed stream, which is not cooled and not partially condensed, is in the range of 95:5 to 5:95, or in the range of 95:5 to 65:35, or in the range of 95:5 to 70:30.

A plant for performing some embodiments is shown schematically in FIG. 1. It should be understood that the operating parameters for the plant, such as the temperature, pressure, flow rates and compositions of the various streams, are established to achieve the desired separation and recovery of the NGLs. The required operating parameters also depend on the composition of the feed gas. The required operating parameters can be readily determined by those skilled in the art using known techniques, including for example computer simulations. Accordingly, the descriptions and ranges of the various operating parameters provided below are intended to provide a description of specific embodiments, and they are not intended to limit the scope of the disclosure in any way.

Feed gas is fed through line (12) to main heat exchanger (10). The feed gas may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. The feed gas is typically fed to the main heat exchanger at a temperature between about 110.degree. F. and 130.degree. F. and at a pressure between about 100 psia and 450 psia. The feed gas is cooled and partially liquefied in the main heat exchanger (10) by making heat exchange contact with cooler process streams and with a refrigerant which may be fed to the main heat exchanger through line (15) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane may be used to provide the necessary cooling for the feed gas. The feed gas is cooled in the main heat exchanger to a temperature between about 0.degree. F. and −40.degree. F.

The cool feed gas (12) exits the main heat exchanger (10) and enters the distillation column (20) through feed line (13). The distillation column operates at a pressure slightly below the pressure of the feed gas, typically at a pressure of between about 5 psi and 10 psi less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as for example propane and other C₃+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (16), while the lighter components exit through vapor overhead line (14). Preferably, the bottoms stream (16) exits the distillation column at a temperature of between about 150.degree. F. and 300.degree. F., and the overhead stream (14) exits the distillation column at a temperature of between about −10.degree. F. and −80.degree. F.

The bottoms stream (16) from the distillation column is split, with a product stream (18) and a recycle stream (22) directed to a reboiler (30) which receives heat input (Q). Optionally, the product stream (18) may be cooled in a cooler to a temperature between about 60.degree. F. and 130.degree. F. The product stream (18) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 1, the product stream may highly enriched in propane and heavier components, and ethane and lighter gases are removed as sales gas as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas. The recycle stream (22) is heated in reboiler (30) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.

The distillation column overhead stream (14) passes through main heat exchanger (10), where it is cooled by heat exchange contact with process gases to partially liquefy the stream. The distillation column overhead stream exits the main heat exchanger through line (19) and is cooled sufficiently to produce the mixed refrigerant as described below. Preferably, the distillation column overhead stream is cooled to between about −30.degree. F. and −130.degree. F. in the main heat exchanger.

In the embodiment of the process shown in FIG. 1, the cooled and partially liquefied stream (19) is combined with the overhead stream (28) from reflux separator (40) in mixer (100) and is then fed through line (32) to the distillation column overhead separator (60). Alternatively, stream (19) may be fed to the distillation column overhead separator (60) without being combined with the overhead stream (28) from reflux separator (40). Overhead stream (28) may be fed to the distillation column overhead separator directly, or in other embodiments of the process, the overhead stream (28) from reflux separator (40) may be combined with the sales gas (42). Optionally, the overhead stream from reflux separator (40) may be fed through control valve (75) prior to being fed through line (28 a) to be mixed with distillation column overhead stream (19). Depending upon the feed gas used and other process parameters, control valve (75) may be used to hold pressure on the ethane compressor (80), which can ease condensing this stream and to provide pressure to transfer liquid to the top of the distillation column. Alternatively, a reflux pump can be used to provide the necessary pressure to transfer the liquid to the top of the column.

In the embodiment shown in FIG. 1, the combined distillation column overhead stream and reflux drum overhead stream (32) is separated in the distillation column overhead separator (60) into an overhead stream (42) and a bottoms stream (34). The overhead stream (42) from the distillation column overhead separator (60) contains product sales gas (e.g. methane, ethane and lighter components). The bottoms stream (34) from the distillation column overhead separator is the liquid mixed refrigerant used for cooling in the main heat exchanger (10).

The sales gas flows through the main heat exchanger (10) through line (42) and is warmed. In a typical plant, the sales gas exits the deethanizer overhead separator at a temperature of between about −40.degree. F. and −120.degree. F. and a pressure of between about 85 psia and 435 psia, and exits the main heat exchanger at a temperature of between about 100.degree. F. and 120.degree. F. The sales gas is sent for further processing through line (43).

The mixed refrigerant flows through the distillation column overhead separator bottoms line (34). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (65). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (10). The mixed refrigerant is fed to the main heat exchanger through line (35). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about −60.degree. F. to −175.degree. F. Where the control valve (65) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by between about 20.degree. F. to 50.degree. F. and the pressure is reduced by between about 90 psi to 250 psi. The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger (10) and exits through line (35 a). The temperature of the mixed refrigerant exiting the main heat exchanger is between about 80.degree. F. and 100.degree. F.

After exiting the main heat exchanger, the mixed refrigerant is fed to ethane compressor (80). The mixed refrigerant is compressed to a pressure about 15 psi to 25 psi greater than the operating pressure of the distillation column at a temperature of between about 230.degree. F. to 350.degree. F. By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (36) to cooler (90) where it is cooled to a temperature of between about 70.degree. F. and 130.degree. F. Optionally, cooler (90) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (10) as described below. The compressed mixed refrigerant then flows through line (38) through the main heat exchanger (10) where it is further cooled and partially liquefied. The mixed refrigerant is cooled in the main heat exchanger to a temperature of between about 15.degree. F. to −70.degree. F. The partially liquefied mixed refrigerant is introduced through line (39) to the reflux separator (40). As described previously, in the embodiment of FIG. 1, the overhead (28) from reflux separator (40) is combined with the overheads (14) from the distillation column and the combined stream (32) is fed to the distillation column overhead separator. The liquid bottoms (26) from the reflux separator (40) are fed back to the distillation column as a reflux stream (26). Control valves (75, 85) may be used to hold pressure on the compressor to promote condensation.

The open loop mixed refrigerant used as reflux enriches the distillation column with gas phase components. With the gas in the distillation column enriched, the overhead stream of the column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than normally required for high recovery of NGLs.

The reflux to the distillation column also reduces losses of heavier hydrocarbons from the column. For example, in processes for recovery of propane, the reflux increases the mole fraction of ethane in the distillation column, which makes it easier to condense the overhead stream. The process uses the liquid condensed in the distillation column overhead drum twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column.

In another embodiment shown in FIG. 2, in which like numbers indicate like components and flow streams described above, the process is used to separate propane and other C₃+ hydrocarbons from ethane and light components. A tee (110) is provided in line (38) after the mixed refrigerant compressor (80) and the mixed refrigerant cooler to split the mixed refrigerant into a return line (45) and an ethane recovery line (47). The return line (45) returns a portion of the mixed refrigerant to the process through main heat exchanger (10) as described above. Ethane recovery line (47) supplies a portion of the mixed refrigerant to a separate ethane recovery unit for ethane recovery. Removal of a portion of the mixed refrigerant stream has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used, for example, as a feed stream in an ethylene cracking unit.

In another embodiment, the NGL recovery unit can recover significant amounts of ethane with the propane. In this embodiment of the process, the distillation column is a demethanizer, and the overhead stream contains primarily methane and inert gases, while the column bottoms contain ethane, propane and heavier components.

In another embodiment of the process, the deethanizer overhead drum may be replaced by an absorber. As shown in FIG. 3, in which like numbers indicate like components and flow streams described above, in this embodiment, the overhead stream (14) from the distillation column (20) passes through main heat exchanger (10) and the cooled stream (19) is fed to absorber (120). The overhead stream (28) from reflux separator (40) is also fed to the absorber (120). The overhead stream (42) from the absorber is the sales gas and the bottoms stream (34) from the absorber is the mixed refrigerant. The other streams and components shown in FIG. 3 have the same flow paths as described above.

In yet another embodiment shown in FIG. 4, in which like numbers indicate like components and flow streams described above, the second separator and the cooler are not used in the process. In this embodiment, the compressed mixed refrigerant (36) is fed through the main heat exchanger (10) and fed to the distillation tower through line (39) to provide reflux flow.

In the embodiment shown in FIG. 5, the gas feed stream (112) is split to create a first feed stream (112 a) and a second feed stream (112 b). The first feed stream (112 a) enters the heat exchanger (110) for cooling to form a cold feed stream (113) from the heat exchanger (110) that is partially liquefied to form a stream (113) containing a mixture of liquid and vapor. The second feed stream (112 b) is a warm gas by-pass feed stream that is not pre-cooled and typically is entirely in a gas phase, with no liquid. A valve (195) is provided for the second feed stream (112 b) for process control purposes, including controlling the relative flow rates of the first and second feed streams (112 a) and (112 b) into the distillation column (120). When liquids condense in the first feed stream (112 a), a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (112 b) has the same overall composition as first feed stream (112 a) and the cooled feed stream (113) but typically does not contain any liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of the second feed stream (112 b) than in the cold feed stream (113). Feeding the warm by-pass gas of the second feed stream (112 b) to the distillation column one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (113) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (120). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include second feed stream (112 b).

The feed gas (112) may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. The feed gas in the first feed stream (112 a) is typically fed to the main heat exchanger at a temperature between about 110.degree. F. and 130.degree. F. and at a pressure between about 100 psia and 450 psia. The feed as is cooled and partially liquefied in the main heat exchanger (110) by making heat exchange contact with cooler process streams and with a refrigerant which may be fed to the main heat exchanger through line (115) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane may be used to provide the necessary cooling for the feed gas. The feed gas is cooled in the main heat exchanger to a temperature between about 0.degree. F. and −40.degree. F.

The cool feed gas (112 a) exits the main heat exchanger (110) and enters the distillation column (120) through feed line (113). The distillation column operates at a pressure slightly below the pressure of the feed gas, typically at a pressure of between about 5 psi and 10 psi less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as for example propane and other C₃+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (116), while the lighter components exit through vapor overhead line (114). Preferably, the bottoms stream (116) exits the distillation column at a temperature of between about 150.degree. F. and 300.degree. F., and the overhead stream (114) exits the distillation column at a temperature of between about −10.degree. F. and −80.degree. F.

The bottoms stream (116) from the distillation column is split, with a product stream (118) and a recycle stream (122) directed to a reboiler (130) which receives heat input (Q). Optionally, the product stream (118) may be cooled in a cooler to a temperature between about 60.degree. F. and 130.degree. F. The product stream (118) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 5, the product stream may be highly enriched in propane and heavier components, and ethane and lighter gases are removed as sales gas in the sales gas line (143) as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas. The recycle stream (122) is heated in reboiler (130) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.

The distillation column overhead stream (114) passes through main heat exchanger (110), where it is cooled by heat exchange contact with process gases to partially liquefy the stream. The distillation column overhead stream exits the main heat exchanger through line (119) and is cooled sufficiently to produce the mixed refrigerant as described below. Preferably, the distillation column overhead stream is cooled to between about −30.degree. F. and −130.degree. F. in the main heat exchanger.

In the embodiment of the process shown in FIG. 5, the cooled and partially liquefied stream (119) is combined with the overhead stream (128) from reflux separator (140) in mixer (200) and is then fed through line (132) to the distillation column overhead separator (160). Alternatively, stream (119) may be fed to the distillation column overhead separator (160) without being combined with the overhead stream (128) from reflux separator (140). Overhead stream (128) may be fed to the distillation column overhead separator directly, or in other embodiments of the process, the overhead stream (128) from reflux separator (140) may be combined with the sales gas (142). Optionally, the overhead stream from reflux separator (140) may be fed through control valve (175) prior to being fed through line (128 a) to be mixed with distillation column overhead stream (119). Depending upon the feed gas used and other process parameters, control valve (175) may be used to hold pressure on the ethane compressor (180), which can ease condensing this stream and to provide pressure to transfer liquid to the top of the distillation column. Alternatively, a reflux pump can be used to provide the necessary pressure to transfer the liquid to the top of the column.

In the embodiment shown in FIG. 5, the combined distillation column overhead stream and reflux drum overhead stream (132) is separated in the distillation column overhead separator (160) into an overhead stream (142) and a bottoms stream (134). The overhead stream (142) from the distillation column overhead separator (160) contains product sales gas (e.g. methane, ethane and lighter components). The bottoms stream (134) from the distillation column overhead separator is the liquid mixed refrigerant used for cooling in the main heat exchanger (110).

The sales gas flows through the main heat exchanger (110) through line (142) and is warmed. In a typical plant, the sales gas exits the deethanizer overhead separator at a temperature of between about −40.degree. F. and −120.degree. F. and a pressure of between about 85 psia and 435 psia, and exits the main heat exchanger at a temperature of between about 100.degree. F. and 120.degree. F. The sales gas is sent for further processing through line (143).

The mixed refrigerant flows through the distillation column overhead separator bottoms line (134). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (165). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (110). The mixed refrigerant is fed to the main heat exchanger through line (135). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about −60.degree. F. to −175.degree. F. Where the control valve (165) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by between about 20.degree. F. to 50.degree. F. and the pressure is reduced by between about 90 psi to 250 psi. The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger (110) and exits through line (135 a). The temperature of the mixed refrigerant exiting the main heat exchanger is between about 80.degree. F. and 100.degree. F.

After exiting the main heat exchanger, the mixed refrigerant is fed to ethane compressor (180). The mixed refrigerant is compressed to a pressure about 15 psi to 25 psi greater than the operating pressure of the distillation column at a temperature of between about 230.degree. F. to 350.degree. F. By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (136) to cooler (190) where it is cooled to a temperature of between about 70.degree. F. and 130.degree. F. Optionally, cooler (190) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (110) as described below. The compressed mixed refrigerant then flows through line (138) through the main heat exchanger (110) where it is further cooled and partially liquefied. The mixed refrigerant is cooled in the main heat exchanger to a temperature of between about 15.degree. F. to −70.degree. F. The partially liquefied mixed refrigerant is introduced through line (139) to the reflux separator (140). As described previously, in the embodiment of FIG. 5, the overhead (128) from reflux separator (140) is combined with the overheads (114) from the distillation column and the combined stream (132) is fed to the distillation column overhead separator. The liquid bottoms (126) from the reflux separator (140) are fed back to the distillation column as a reflux stream (126). Control valves (175, 185) may be used to hold pressure on the compressor to promote condensation.

In the embodiment shown in FIG. 6, the gas feed stream (212) is split to create a first feed stream (212 a) and a second feed stream (212 b). The first feed stream (212 a) enters the heat exchanger (210) for cooling to form a cold feed stream (213) from the heat exchanger (210) that is partially liquefied. The second feed stream (212 b) is a warm gas by-pass feed stream that is not pre-cooled and typically is in an entirely gas phase, with no liquid. A valve (295) is provided for the second feed stream (212 b) for process control purposes. When liquids condense, a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (212 b) has the same composition as the first feed stream (212 a) but contains less liquid (and typically is entirely in the gas phase). As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (212 b) than in the cold feed stream (213). Placing the warm by-pass gas of the second feed stream (212 b) one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (213) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (220). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (212 b).

In the embodiment of FIG. 6, the process is used to separate propane and other C₃+ hydrocarbons from ethane and light components. A tee (310) is provided in line (238) after the mixed refrigerant compressor (280) and the mixed refrigerant cooler to split the mixed refrigerant into a return line (245) and an ethane recovery line (247). The return line (245) returns a portion of the mixed refrigerant to the process through main heat exchanger (210) as described above. Ethane recovery line (247) supplies a portion of the mixed refrigerant to a separate ethane recovery unit for ethane recovery. Removal of a portion of the mixed refrigerant stream has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used, for example, as a feed stream in an ethylene cracking unit.

In the embodiment shown in FIG. 7, the gas feed stream (312) is split to create a first feed stream (312 a) that enters the heat exchanger (310) for cooling to form a cold feed stream (313) from the heat exchanger (310) that is partially liquefied, and a second feed stream (312 b) that is a warm gas by-pass feed stream that is not pre-cooled. A valve (395) is provided for the second feed stream (312 b) for process control purposes. When liquids condense in first feed stream (312 a), a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (312 b) has the same composition as the first feed stream (312 a) but contains less liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (312 b) than in the cold feed stream (313). Placing the warm by-pass gas of the second feed stream (312 b) one or more, or 1 to 10, or 1 to 7, or 1 to 4, vapor-liquid equilibrium stages below the cold feed stream (313) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (320). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (312 b).

As is shown in FIG. 7, the deethanizer overhead drum may be replaced by an absorber. In this embodiment, the overhead stream (314) from the distillation column (320) passes through main heat exchanger (310) and the cooled stream (319) is fed to an absorber (321). The overhead stream (328) from reflux separator (340) is also fed to the absorber (321) through line (332). The overhead stream (342) from the absorber (321) is the sales gas and the bottoms stream (334) from the absorber (321) is the mixed refrigerant. The other streams and components shown in FIG. 7 have the same flow paths as described above.

In yet another embodiment shown in FIG. 8 the second separator and the cooler are not used in the process. In this embodiment, the compressed mixed refrigerant (436) is fed through the main heat exchanger (410) and fed to the distillation column (420) through line (439) to provide reflux flow.

In the embodiment shown in FIG. 8, the gas feed stream (412) is split to create a first feed stream (412 a) that enters the heat exchanger (410) for cooling to form a cold feed stream (413) from the heat exchanger (410) that is partially liquefied, and a second feed stream (412 b) that is a warm gas by-pass feed stream that is not pre-cooled. A valve (495) is provided for the second feed stream (412 b) for process control purposes. When liquids condense, a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (412 b) has the same composition as the first feed stream (412 a) but contains less liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (412 b) than in the cold feed stream (413). Placing the warm by-pass gas of the second feed stream (412 b) one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (413) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (420). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (412 b).

In yet another embodiment shown in FIG. 9, the split feed scheme is incorporated into a system that is somewhat similar to a process described in U.S. Pat. No. 8,627,681, the contents of which are incorporated herein by reference in their entirety. The benefits of the embodiment of FIG. 9 are surprisingly and unexpectedly discovered that there will be a decrease in refrigeration duty specification, decrease in deethanizer reboiler duty specification, decrease in deethanizer vapor and liquid traffic thus providing for a distillation column sizing decrease, and a decrease in the refrigeration and reboiler duty specification with high pressure feeds. The total propane and mixed refrigerant compressor duty is over 11 percent higher without the split feed. As is shown, considerable economic benefits from reduced total invested cost and operational costs can be obtained as a result of these unexpected improvements.

More specifically, the overall process of FIG. 9 is designated as 502. Feed stream (512) is split to create first feed stream (512 a) and second feed stream (512 b). First feed stream (512 a) enters the heat exchanger (510) for cooling to form a cold or high pressure stream (513) from the heat exchanger (510) that is partially liquefied. Warm vapor by-pass stream (512 b) is a second stream that is not pre-cooled. Stream (512 b) passes through control valve (605) to reduce its pressure and is then fed to the middle of a distillation column (520) at a location that is one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the entry point of stream 513.

Although a multi-pass heat exchanger (510) is illustrated, use of multiple heat exchangers may be used to achieve similar results, as is also the case with the embodiments shown in FIGS. 5-8. The feed stream (512) may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. In embodiments, the first feed stream (512 a) is typically fed to the main heat exchanger at a temperature between about 43.degree. C. and 54.degree. C. (110.degree. F. and 130.degree. F.) and at a pressure between about 7 bar and 31 bar (100 psia and 450 psia). The first feed stream (512 a) is cooled and partially liquefied in the main heat exchanger 510 via indirect heat exchange with cooler process streams and/or with a refrigerant which may be fed to the main heat exchanger via line (515) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane, for example, may be used to provide the necessary cooling for the feed gas. The feed gas may be cooled in the main heat exchanger to a temperature between about −18.degree. C. and −40.degree. C. (0.degree. F. and −40.degree. F.).

The cool feed gas exits the main heat exchanger (510) and is fed to distillation column (520) via feed line (513). Distillation column (520) operates at a pressure slightly below the pressure of the feed gas, typically at a pressure about 0.3 to 0.7 bar (5 to 10 psi) less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as propane and other C.sub.3+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (516), while the lighter components exit through vapor overhead line (514). In embodiments, the bottoms stream 516 exits the distillation column at a temperature between about 65.degree. C. and 149.degree. C. (150.degree. F. and 300.degree. F.), and the overhead stream 14 exits the distillation column at a temperature of between about −23.degree. C. and −62.degree. C. (−10.degree. F. and −80.degree. F.).

The bottoms stream (516) from the distillation column is split, with a product stream (518) and a reboil stream (522) directed to a reboiler (530). Optionally, the product stream (518) may be cooled in a cooler (not shown) to a temperature between about 515.degree. C. and 554.degree. C. (60.degree. F. and 130.degree. F.). The product stream (518) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 9, the product stream may be enriched in propane and heavier components, and ethane and lighter gases are further processed as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas produced. The reboil stream (522) is heated in reboiler (530) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.

The distillation column overhead stream (514) passes through main heat exchanger (510), where it is cooled by indirect heat exchange with process gases to at least partially liquefy or completely (100%) liquefy the stream. The distillation column overhead stream exits the main heat exchanger (510) through line (519) and is cooled sufficiently to produce the mixed refrigerant as described below. In some embodiments, the distillation column overhead stream is cooled to between about −34.degree. C. and −90.degree. C. (−30.degree. F. and −130.degree. F.) in main heat exchanger 510.

The cooled and partially liquefied stream (519) and the overhead stream (528) (stream 532 following control valve 575) from reflux separator (540) may be fed to distillation column overhead separator (585).

The components in distillation column overhead stream (519) and reflux drum overhead stream (532) are separated in overhead separator (585) into an overhead stream (542), a side draw fraction (551), and a bottoms stream (534). The overhead stream (542) from distillation column overhead separator (585) contains methane, ethane, nitrogen, and other lighter components, and is enriched in nitrogen content. Side draw fraction (551) may be of intermediate nitrogen content. The bottoms stream (534) from distillation column overhead separator (585) is the liquid mixed refrigerant used for cooling in the main heat exchanger (510), which may be depleted in nitrogen content. The side draw fraction may be reduced in pressure across flow valve (595), fed to heat exchanger (510) for use in the integrated heat exchange system, and recovered via flow line (552).

The components in overhead stream (542) are fed to main heat exchanger (510) and warmed. In a typical plant, the overhead fraction recovered via stream (542) from overhead separator (585) is at a temperature between about −40.degree. C. and −84.degree. C. (−40.degree. F. and −120.degree. F.) and at a pressure between about 5 bar and 30 bar (85 psia and 435 psia). Following heat exchange in main heat exchanger (510), the overhead fraction recovered from heat exchanger 510 via stream (543) may be at a temperature between about 37.degree. C. and 49.degree. C. (100.degree. F. and 120.degree. F.). The overhead fraction is enriched in nitrogen content and may be recovered via stream (543) as a low-btu natural gas stream.

The mixed refrigerant, as mentioned above, is recovered from distillation column overhead separator (585) via bottoms line (534). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (565). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (510). The mixed refrigerant is fed to the main heat exchanger through line (535). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about −51.degree. C. and −115.degree. C. (−60.degree. F. to −175.degree. F.).

Where the control valve (565) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by about 6.degree. C. to 10.degree. C. (20.degree. F. to 50.degree. F.) and the pressure is reduced by about 6 bar to 17 bar (90 to 250 psi). The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger 510 and exits through line (535 a).

The temperature of the mixed refrigerant exiting the main heat exchanger is between about 26.degree. C. and 38.degree. C. (80.degree. F. and 100.degree. F.).

After exiting main heat exchanger (510), the mixed refrigerant is fed to compressor (580). The mixed refrigerant is compressed to a pressure 1 bar to 2 bar (15 psi to 25 psi) greater than the operating pressure of the distillation column, and at a temperature between about 110.degree. C. to 177.degree. C. (230.degree. F. to 350.degree. F.). By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (536) to cooler (590) where it is cooled to a temperature between about 21.degree. C. and 54.degree. C. (70.degree. F. and 130.degree. F.). Optionally, cooler (590) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (510). The compressed mixed refrigerant then flows via line (538) through the main heat exchanger (510) where it is further cooled and partially liquefied.

The mixed refrigerant is cooled in the main heat exchanger to a temperature from about −9.degree. C. to −57.degree. C. (15.degree. F. to −70.degree. F.) The partially liquefied mixed refrigerant is introduced through line (539) to reflux separator (540). As described previously, the overheads (528) from reflux separator (540) and overheads (514) from the distillation column (520) are fed to the distillation column overhead separator (585). The liquid bottoms (526) from the reflux separator (540) are fed back to the distillation column (520) as a reflux stream (526). Control valves (575), (586) may be used to hold pressure on the compressor to promote condensation.

The mixed refrigerant used as reflux (fed via stream 526) enriches distillation column (520) with gas phase components. With the gas in the distillation column enriched, the overhead stream of the column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than normally required for a high recovery of NGLs.

The reflux to distillation column (520) also reduces heavier hydrocarbons in the overheads fraction. For example, in processes for recovery of propane, the reflux increases the mole fraction of ethane in the distillation column, which makes it easier to condense the overhead stream. The process uses the liquid condensed in the distillation column overhead separator twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column.

At least a portion of the mixed refrigerant in flow line (528), having a very low nitrogen content, may be withdrawn via flow stream (532 ex) prior to separator (585). In some embodiments, the portion withdrawn via flow stream (532 ex) may be used for pipeline sales. In other embodiments, a mixed refrigerant stream (532 ex), having less than 1 mole % nitrogen, may be mixed with a high or intermediate btu natural gas process stream having greater than 4% nitrogen to result in a pipeline sales stream having 4% or less nitrogen.

For example, mixed refrigerant stream (532 ex) may be combined with intermediate btu natural gas in stream (551) (side draw) to result in a natural gas stream suitable for pipeline sales. The flow rates of streams (532 ex) and (551) may be such that the resulting product stream (548) has a nitrogen (inert) content of less than 4 mole %. In some embodiments, flow stream (532 ex) may be fed to main heat exchanger (510); and following heat transfer, the mixed refrigerant may be recovered from heat exchanger (510) via flow line (541) for admixture with intermediate btu stream (551). Other process streams may also be admixed with mixed refrigerant stream (532 ex) in other embodiments.

Processes according to the embodiment of FIG. 9 allows for substantial process flexibility, providing for the ability to efficiently process feed gas streams having a wide range of nitrogen content. The embodiment described with regard to FIG. 9 allows for recovery of a majority of the feed gas btu value as a natural gas sales stream. Iso-pressure open refrigeration processes according to embodiments disclosed herein may additionally include separation of nitrogen from high or intermediate nitrogen content streams, allowing for additional recovery of btu value or additional flexibility with regard to process conditions and feed gas nitrogen content.

Examples of specific embodiments of the processes are described below. These examples are provided to further describe the processes described herein and they are not intended to limit the full scope of the disclosure in any way.

CONTROL EXAMPLE 1

In the following examples, operation of the processing plant shown in FIG. 1 with different types and compositions of feed gas were computer simulated using process the Apsen HYSYS simulator. In this example, the operating parameters for C₃+ recovery using a relatively lean feed gas are provided. Table 1 shows the operating parameters for propane recovery using a lean feed gas. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 2. Energy inputs for this embodiment included about 3.717×10⁵5 Btu/hr (Q) to the reboiler (30) and about 459 horsepower (P) to the ethane compressor (80).

As can be seen in Table 2, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. Approximately 99.6% of the propane in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of methane and ethane, but contains more propane than the sales gas.

CONTROL EXAMPLE 2

In this example, operating parameters are provided for the processing plant shown in

FIG. 1 using a refinery feed gas for recovery of C₃+ components in the product stream. Table 3 shows the operating parameters using the refinery feed gas. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 4. Energy inputs for this embodiment included about 2.205×10⁶ Btu/hr (Q) to the reboiler (30) and about 228 horsepower (P) to the ethane compressor (80).

As can be seen in Table 4, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases, in particular hydrogen. This stream could be used to feed a membrane unit or PSA to upgrade this stream to useful hydrogen. Approximately 97.2% of the propane in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of methane and ethane, but contains more propane than the sales gas.

CONTROL EXAMPLE 3

In this example, operating parameters are provided for the processing plant shown in

FIG. 1 using a refinery feed gas for the recovery of C.sub.4+ components in the product stream, with the C₃ components removed in the sales gas stream. Table 5 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C.sub.4+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 6. Energy inputs for this embodiment included about 2.512×10⁶ Btu/hr (Q) to the reboiler (30) and about 198 horsepower (P) to the ethane compressor (80).

As can be seen in Table 6, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C.sub.4+ components, while the sales gas stream (43) contains almost entirely C₃ and lighter hydrocarbons and gases. Approximately 99.7% of the C.sub.4+ components in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of C₃ and lighter components, but contains more butane than the sales gas.

CONTROL EXAMPLE 4

In this example, operating parameters are provided for the processing plant shown in FIG. 2 using a refinery feed gas for recovery of C₃+ components in the product stream, with the C₂ and lighter components removed in the sales gas stream. In this embodiment, a portion of the mixed refrigerant is removed through line (47) and fed to an ethane recovery unit for further processing. Table 7 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 8. Energy inputs for this embodiment included about 2.089×10⁶ Btu/hr (Q) to the reboiler (30) and about 391 horsepower (P) to the ethane compressor (80).

As can be seen in Table 8, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

CONTROL EXAMPLE 5

In this example, operating parameters are provided for the processing plant shown in

FIG. 3 using a lean feed gas for recovery of C₃+ components in the product stream, with the C₂ and lighter components removed in the sales gas stream. In this embodiment, an absorber (120) is used to separate the distillation column overhead stream and the reflux separator overhead stream to obtain the mixed refrigerant. Table 9 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 10. Energy inputs for this embodiment included about 3.734×10⁵ Btu/hr (Q) to the reboiler (30) and about 316 horsepower (P) to the ethane compressor (80).

As can be seen in Table 10, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

CONTROL EXAMPLE 6

In this example, operating parameters are provided for the processing plant shown in FIG. 1 using a rich feed gas for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. Table 11 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 12. Energy inputs for this embodiment included about 1.458×10⁶ Btu/hr (Q) to the reboiler (30) and about 226 horsepower (P) to the ethane compressor (80).

As can be seen in Table 12, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

EXAMPLE 7

In this example, operating parameters comparable to the prior examples are provided for a simulated processing plant shown in FIG. 5 using the rich feed gas of Control Example 6 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. Energy inputs for this embodiment included about 1.117×10⁶ Btu/hr (Q) to the reboiler (130) and a reduced horsepower to the ethane compressor (180). In this embodiment, about 15 weight % of the gas feed stream (112) formed the bypass stream (112 b) and the remainder of stream (112) formed the first feed stream (112 a).

As was the case in the prior examples, in this embodiment, the product stream (118) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (143) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas. Based on process simulation data, with the insertion of the feed stream (112 b) into the column (120) as shown, it was discovered that there was an unexpected and surprising decrease in refrigeration duty specification, decrease in deethanizer reboiler duty specification, decrease in deethanizer vapor and liquid traffic thus providing for a distillation column sizing decrease, and a decrease in the refrigeration and reboiler duty specification with high pressure feeds. The total propane and mixed refrigerant compressor duty is over 12 percent higher without the split feed. This results in significant economic savings in both total invested cost (TIC) in the plant and operational costs. By way of illustration, in the USA, 200 MMSCFD gas plants are fairly typical. Such a plant may have about 15,000 HP of refrigeration compressors, depending of feed composition. Based on the following calculation, configuration results in about a 12 percent saving in compressor duty. 15,000 HP×0.12 Savings×0.746 kw/hp×0.1 $/kwh=134 $/hr, which is over a million dollars per year of electric power.

For a 200 MMSCFD plant, a reboiler duty without the split feed is approximately 29.2 MMBTU/HR. In this Example, the reboiler duty is about 22.2 MMBTU/HR. Assuming an energy cost of US 5.00 MMBTU, annual savings would be about $307,000.

EXAMPLE 8

In this set of examples, operating parameters comparable to the prior examples were provided for a simulated processing plant shown in FIG. 5 using the same lean feed gas and product stream compositions as were used in Control Example 1 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass feed stream contained about 10-15 weight % of the feed gas stream 112. Energy inputs for this embodiment were about 20-27% lower than the energy input for Control Example 1. This set of examples resulted in significant economic savings in both total invested cost (TIC) in the plant and operational costs.

PROPHETIC EXAMPLE 9

In this example, operating parameters comparable to the prior examples are provided for a simulated processing plant shown in FIG. 6 using the same lean feed gas and product stream compositions as were used in Control Example 4 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass feed stream contains about 10-15 weight % of the feed gas stream 212. Energy inputs for this embodiment are lower than the energy input for Control Example 4. This embodiment results in significant economic savings in both total invested cost (TIC) in the plant and operational costs.

PROPHETIC EXAMPLE 10

In this example, operating parameters comparable to the prior examples are provided for a simulated processing plant shown in FIG. 7 using the same lean feed gas and product stream compositions as were used in Control Example 5 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass stream contains about 10-15 weight % of the feed gas stream 312. Energy inputs for this embodiment are lower than the energy input for Control Example 5. This embodiment results in significant economic savings in both total invested cost (TIC) in the plant and operational costs.

While specific embodiments have been described above, one skilled in the art will recognize that numerous variations or changes may be made to the process described above without departing from the scope as recited in the appended claims. Accordingly, the foregoing description of preferred embodiments is intended to describe the embodiments an exemplary, rather than a limiting, sense.

TABLE 1 Material Streams 12 13 19 15 17 Vapour 1.0000 0.9838 0.3989 0.0000 0.5000 Fraction Temperature F. 120.0 −25.00 −129.0 −30.00 −29.68 Pressure psia 415.0 410.0 400.0 21.88 20.88 Molar Flow MMSCFD 10.00 10.00 11.76 1.317 1.317 Mass Flow lb/hr 1.973e+004 1.973e+004 2.362e+004 6356 6356 Liquid barrel/day 4203 4203 5100 862.2 862.2 Volume Flow 14 18 32 34 42 Vapour 1.0000 0.0000 0.6145 0.0000 1.0000 Fraction Temperature F. −76.88 251.9 −118.6 −118.7 −118.7 Pressure psia 405.0 410.0 400.0 400.0 400.0 Molar Flow MMSCFD 11.76 0.2517 15.89 6.139 9.723 Mass Flow lb/hr 2.362e+004 1671 3.220e+004 1.414e+004 1.800e+004 Liquid barrel/day 5100 196.3 6931 2925 3995 Volume Flow 43 35 35a 36 38 Vapour 1.0000 0.2758 1.0000 1.0000 1.0000 Fraction Temperature F. 110.0 −165.0 90.00 262.2 120.0 Pressure psia 395.0 149.9 144.9 470.0 465.0 Molar Flow MMSCFD 9.723 6.139 6.139 6.139 6.139 Mass Flow lb/hr 1.800e+004 1.414e+004 1.414e+004 1.414e+004 1.414e+004 Liquid barrel/day 3995 2925 2925 2925 2925 Volume Flow 39 28 26 26a 28a Vapour 0.6723 1.0000 0.0000 0.0452 .09925 Fraction Temperature F. −63.00 −63.00 −63.00 −68.04 −69.27 Pressure psia 460.0 460.0 460.0 415.0 400.0 Molar Flow MMSCFD 6.139 4.127 2.011 2.011 4.127 Mass Flow lb/hr 1.414e+004 8573 5566 5566 8573 Liquid barrel/day 2925 1831 1094 1094 1831 Volume Flow

TABLE 2 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12) (18) (43) Refrigerant (35) Methane 0.9212 0.0000 0.9453 0.6671 Ethane 0.0396 0.0082 0.0402 0.3121 Propane 0.0105 0.4116 0.0001 0.0046 Butane 0.0036 0.1430 0.0000 0.0000 Pentane 0.0090 0.3576 0.0000 0.0000 Heptane 0.0020 0.0795 0.0000 0.0000 CO₂ 0.0050 0.0000 0.0051 0.0145 Nitrogen 0.0091 0.0000 0.0094 0.0017

TABLE 3 Material Streams 12 13 19 15 17 Vapour 0.9617 0.7601 0.7649 0.0000 0.5000 Fraction Temperature F. 120.0 −5.00 −85.00 −15.00 −14.37 Pressure psia 200.0 195.0 185.0 30.12 29.12 Molar Flow MMSCFD 10.00 10.00 9.821 8.498 8.498 Mass Flow lb/hr 2.673e+004 2.673e+004 1.852e+004 4.102e+004 4.102e+004 Liquid barrel/day 4723 4723 4252 5564 5564 Volume Flow 14 18 32 34 42 Vapour 1.0000 0.0000 0.7669 0.0000 1.0000 Fraction Temperature F. −50.25 162.6 −84.09 −84.07 −84.07 Pressure psia 190.0 195.0 185.0 185.0 185.0 Molar Flow MMSCFD 9.821 2.377 9.937 2.314 7.617 Mass Flow lb/hr 1.852e+004 1.559e+004 1.883e+004 7696 1.112e+004 Liquid barrel/day 4252 1844 4314 1436 2876 Volume Flow 43 35 35a 36 38 Vapour 1.0000 0.0833 1.0000 1.0000 1.0000 Fraction Temperature F. 110.0 −103.0 90.00 260.4 120.0 Pressure psia 180.0 50.8 45.8 215.0 210.0 Molar Flow MMSCFD 7.617 2.314 2.314 2.314 2.314 Mass Flow lb/hr 1.112e+004 7696 7696 7696 7696 Liquid barrel/day 2876 1436 1436 1436 1436 Volume Flow 39 28 26 26a 28a Vapour 0.0500 1.0000 0.0000 0.0032 1.0000 Fraction Temperature F. −29.77 −29.77 −29.77 −30.32 −33.30 Pressure psia 205.0 205.0 205.0 200.0 185.0 Molar Flow MMSCFD 2.314 0.1157 2.198 2.198 0.1157 Mass Flow lb/hr 7696 308.1 7388 7388 308.1 Liquid barrel/day 1436 62.34 1373 1373 62.34 Volume Flow

TABLE 4 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12) (18) (43) Refrigerant (35) Hydrogen 0.3401 0.0000 0.4465 0.0038 Methane 0.2334 0.0000 0.3062 0.0658 Ethane 0.1887 0.0100 0.2439 0.8415 Propane 0.0924 0.3783 0.0034 0.0889 Butane 0.0769 0.3234 0.0000 0.0000 Pentane 0.0419 0.1760 0.0000 0.0000 Heptane 0.0267 0.1124 0.0000 0.0000 CO₂ 0.0000 0.0000 0.0000 0.0000 Nitrogen 0.0000 0.0000 0.0000 0.0000

TABLE 5 Material Streams 12 13 19 15 17 Vapour 0.9805 0.8125 0.8225 0.0000 0.5000 Fraction Temperature F. 120.0 0.00 −43.00 −20.00 −19.46 Pressure psia 135.0 130.0 120.0 27.15 26.15 Molar Flow MMSCFD 10.00 10.00 10.31 8.058 8.058 Mass Flow lb/hr 2.673e+004 2.673e+004 2.339e+004 3.890e+004 3.890e+004 Liquid barrel/day 4723 4723 4624 5276 5276 Volume Flow 14 18 32 34 42 Vapour 1.0000 0.0000 0.8234 0.0000 1.0000 Fraction Temperature F. −13.13 195.3 −42.52 −42.49 −42.49 Pressure psia 125.0 130.0 120.0 120.0 120.0 Molar Flow MMSCFD 10.31 1.462 10.38 1.840 8.557 Mass Flow lb/hr 2.339e+004 1.119e+004 2.360e+004 8068 1.561e+004 Liquid barrel/day 4624 1245 4661 1183 3490 Volume Flow 43 35 35a 36 38 Vapour 1.0000 0.0805 1.0000 1.0000 1.0000 Fraction Temperature F. 110.0 −62.0 90.00 238.2 120.0 Pressure psia 115.0 31.75 26.75 150.0 145.0 Molar Flow MMSCFD 8.557 1.840 1.840 1.840 1.840 Mass Flow lb/hr 1.561e+004 8068 8068 8068 8068 Liquid barrel/day 3490 1183 1183 1183 1183 Volume Flow 39 28 26 26a 28a Vapour 0.0349 1.0000 0.0000 0.0038 1.0000 Fraction Temperature F. 15.00 15.00 15.00 14.31 11.44 Pressure psia 140.0 140.0 140.0 135.0 120.0 Molar Flow MMSCFD 1.840 6.425e−002 1.776 1.776 6.425e−002 Mass Flow lb/hr 8068 211.4 7856 7856 211.4 Liquid barrel/day 1183 36.58 1147 1147 36.58 Volume Flow

TABLE 6 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12) (18) (43) Refrigerant (35) Hydrogen 0.3401 0.0000 0.3975 0.0022 Methane 0.2334 0.0000 0.2728 0.0257 Ethane 0.1887 0.0000 0.2220 0.2461 Propane 0.0924 0.0100 0.1074 0.7188 Butane 0.0769 0.5212 0.0003 0.0071 Pentane 0.0419 0.2861 0.0000 0.0000 Heptane 0.0267 0.1828 0.0000 0.0000 CO₂ 0.0000 0.0000 0.0000 0.0000 Nitrogen 0.0000 0.0000 0.0000 0.0000

TABLE 7 Material Streams 12 13 19 15 17 14 Vapour 0.9617 0.7202 0.6831 0.0000 0.5000 1.0000 Fraction Temperature F. 120.0 −25.00 −145.0 −30.00 −29.68 −22.80 Pressure psia 200.0 195.0 185.0 21.88 20.88 190.0 Molar Flow MMSCFD 10.00 10.00 8.153 7.268 7.628 8.153 Mass Flow lb/hr 2.673e+004 2.673e+004 1.367e+004 3.508e+004 3.508e+004 1.367e+004 Liquid barrel/day 4723 4723 3231 4758 4758 3231 Volume Flow 18 32 34 42 43 Vapour 0.0000 0.6833 0.0000 1.0000 1.000 Fraction Temperature F. 176.0 −144.9 −144.9 −144.9 110.0 Pressure psia 195.0 185.0 185.0 185.0 180.0 Molar Flow MMSCFD 1.970 8.160 2.589 5.576 5.576 Mass Flow lb/hr 1.348e+004 1.369e+004 8758 4943 4943 Liquid barrel/day 1567 3234 1570 1661 1667 Volume Flow 35 35a 36 38 39 28 Vapour 0.0957 1.0000 1.0000 1.0000 0.0500 1.0000 Fraction Temperature F. −163.1 90.00 330.0 120.0 −61.75 −61.75 Pressure psia 28.00 23.00 215.0 210.0 205.0 205.0 Molar Flow MMSCFD 2.589 2.589 2.589 2.589 0.1294 6.472e−003 Mass Flow lb/hr 8758 8758 8758 8758 437.9 14.05 Liquid barrel/day 1570 1570 1570 1570 78.48 3.009 Volume Flow Vapour 26 26a 28a 45 47 Fraction 0.0000 0.0028 1.0000 1.000 1.0000 Temperature F. Pressure psia −61.75 −62.15 −64.65 120.0 120.0 Molar Flow MMSCFD 205.0 200.0 185.0 210.0 210.0 Mass Flow lb/hr 0.1230 0.1230 6.472e−003 0.1294 2.459 Liquid barrel/day 423.8 423.8 14.05 437.9 8320 Volume 75.47 75.47 3.009 78.48 1491 Flow

TABLE 8 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12) (18) (43) Refrigerant (35) Hydrogen 0.3401 0.0000 0.6085 0.0034 Methane 0.2334 0.0000 0.3517 0.1520 Ethane 0.1887 0.0100 0.0392 0.6719 Propane 0.0924 0.2974 0.0006 0.1363 Butane 0.0769 0.3482 0.0000 0.0335 Pentane 0.0419 0.2087 0.0000 0.0028 Heptane 0.0267 0.1828 0.0000 0.0000 CO₂ 0.0000 0.1357 0.0000 0.0000 Nitrogen 0.0000 0.0000 0.0000 0.0000

TABLE 9 Material Streams 12 13 19 15 17 Vapour 1.0000 0.9838 0.6646 0.0000 0.5000 Fraction Temperature F. 120.0 −25.00 −119.0 −30.00 −29.68 Pressure psia 415.0 410.0 400.0 21.88 20.88 Molar Flow MMSCFD 10.00 10.00 11.83 1.263 1.263 Mass Flow lb/hr 1.973e+004 1.973e+004 2.369e+004 6096 6096 Liquid barrel/day 4203 4203 5115 826.9 826.9 Volume Flow 14 18 32 34 42 Vapour 1.0000 0.0000 0.9925 0.0000 1.0000 Fraction Temperature F. −79.00 251.1 −77.01 −109.5 −118.9 Pressure psia 405.0 410.0 405.0 405.0 400.0 Molar Flow MMSCFD 11.83 0.2534 1.577 3.668 9.730 Mass Flow lb/hr 2.369e+004 1679 3206 8867 1.801e+004 Liquid barrel/day 5115 197.4 688.7 1804 3997 Volume Flow 35 35a 36 38 39 Vapour 0.3049 1.0000 1.0000 1.0000 0.4300 Fraction Temperature F. −162.0 90.00 280.9 120.0 −71.34 Pressure psia 128.30 123.30 470.0 465.0 460.0 Molar Flow MMSCFD 3.668 3.668 3.668 3.668 3.688 Mass Flow lb/hr 8867 8867 8867 8867 8867 Liquid barrel/day 1804 1804 1804 1804 1804 Volume Flow 28 26 26a 43 Vapour 1.0000 0.0000 0.0464 1.000 Fraction Temperature F. −71.34 −71.34 −76.54 110.0 Pressure psia 460.0 460.0 415.0 395.0 Molar Flow MMSCFD 1.577 2.091 2.091 9.730 Mass Flow lb/hr 3206 5661 5661 1.801e+004 Liquid barrel/day 688.7 1115 1115 3997 Volume Flow

TABLE 10 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12) (18) (43) Refrigerant (35) Methane 0.9212 0.0000 0.9457 0.5987 Ethane 0.0396 0.0083 0.0397 0.3763 Propane 0.0105 0.4154 0.0001 0.0054 Butane 00036. 0.1421 0.0000 0.0000 Pentane 0.0090 0.3552 0.0000 0.0000 Heptane 0.0020 0.0789 0.0000 0.0000 CO₂ 0.0050 0.0000 0.0051 0.0195 Nitrogen 0.0091 0.0000 0.0094 0.0001

TABLE 11 Material Streams 12 13 19 15 17 Vapour 1.0000 0.8833 0.7394 0.0000 0.5000 Fraction Temperature F. 120.0 −20.00 −85.5 −30.00 −29.68 Pressure psia 315.0 310.0 305.0 21.88 20.88 Molar Flow MMSCFD 10.00 10.00 11.37 5.018 5.018 Mass Flow lb/hr 2.484e+004 2.484e+004 2.549e+004 2.422e+004 2.422e+004 Liquid barrel/day 4721 4721 5338 3285 3285 Volume Flow 14 18 32 34 42 Vapour 1.0000 0.0000 0.7491 0.0000 1.0000 Fraction Temperature F. −55.13 181.7 −84.23 −84.24 −84.24 Pressure psia 310.0 315.0 305.0 305.0 305.0 Molar Flow MMSCFD 11.37 1.139 11.81 2.952 8.844 Mass Flow lb/hr 2.549e+004 6778 2.648e+004 8419 1.802e+004 Liquid barrel/day 5338 834.5 5546 1660 3877 Volume Flow 43 35 35a 36 38 Vapour 1.0000 0.2044 1.0000 1.0000 1.0000 Fraction Temperature F. 110.0 −120.0 90.00 246.2 120.0 Pressure psia 300.0 113.9 108.9 375.0 370.0 Molar Flow MMSCFD 8.844 2.952 2952 2952 2952 Mass Flow lb/hr 1.802e+004 8419 8419 8419 8419 Liquid barrel/day 3877 1660 1660 1660 1660 Volume Flow 39 28 26 26a 28a Vapour 0.1500 1.0000 0.0000 0.0434 .09975 Fraction Temperature F. −49.05 −49.05 −49.05 −54.73 −57.22 Pressure psia 365.0 365.0 365.0 320.0 305.0 Molar Flow MMSCFD 2952 0.4429 2.510 2.510 0.4429 Mass Flow lb/hr 8419 990.7 7429 7429 990.7 Liquid barrel/day 1660 207.9 1452 1452 207.9 Volume Flow

TABLE 12 Mole Fractions of Components in Streams Feed Gas Product Sales Gas Mixed (12)) (18) (43) Refrigerant (35) Methane 0.7304 0.0000 0.8252 0.3071 Ethane 0.1429 0.0119 0.1566 0.6770 Propane 0.0681 0.5974 0.0003 0.0071 Butane 0.0257 0.2256 0.0000 0.0000 Pentane 0.0088 0.0772 0.0000 0.0000 Heptane 0.0100 0.0878 0.0000 0.0000 CO₂ 0.0050 0.0000 0.0056 0.0079 Nitrogen 0.0091 0.0000 0.0103 0.0009 

What is claimed is:
 1. A process for recovery of natural gas liquids from a feed gas stream, comprising: (a) forming a first portion of the feed gas stream and a second portion of the feed gas stream, wherein the mass ratio of the first portion to the second portion is in the range of 95:5 to 5:95; (b) cooling the first portion in a heat exchanger and at least partially condensing the first portion; (c) feeding the second portion and the cooled and at least partially condensed first portion to a distillation column wherein lighter components are removed from the distillation column as an overhead vapor stream and heavier components are removed from the distillation column in the bottoms as a product stream, and wherein the second portion is fed into the distillation column at a point one or more vapor-liquid equilibrium stages below the first portion, thereby allowing mass transfer exchange between liquids of the cooled first portion and vapors of the second portion within the column; (d) feeding the distillation column overhead stream to the heat exchanger and cooling the distillation column overhead stream to at least partially liquefy the distillation column overhead stream; (e) feeding the at least partially liquefied distillation column overhead stream to a first separator; (f) separating the vapor and liquid in the first separator to produce an overhead vapor stream comprising sales gas and a bottoms stream comprising a mixed refrigerant; (g) feeding the mixed refrigerant stream to the heat exchanger to provide cooling, wherein the mixed refrigerant stream vaporizes as it passes through the heat exchanger; (h) compressing the vaporized mixed refrigerant stream and passing the compressed mixed refrigerant stream through the heat exchanger; and (i) feeding at least a portion of the compressed mixed refrigerant stream to the distillation column as a reflux stream.
 2. The process of claim 1, further comprising, before (i), feeding the compressed mixed refrigerant stream to a second separator, and feeding the bottoms from the second separator to the distillation column as the reflux stream.
 3. The process of claim 1, further comprising reducing the temperature of the mixed refrigerant stream before the mixed refrigerant stream enters the heat exchanger by reducing the pressure of the mixed refrigerant using a control valve.
 4. The process of claim 1, further comprising combining the overhead stream from the second separator with the overhead stream from the distillation column and feeding the combined stream to the first separator.
 5. The process of claim 1, further comprising cooling the compressed mixed refrigerant in a cooler before passing the compressed mixed refrigerant stream through the heat exchanger.
 6. The process of claim 1, wherein the first separator is an absorber.
 7. The process of claim 1, wherein the feed gas stream is one of natural gas or refinery gas.
 8. The process of claim 1, wherein the distillation column operates at a pressure of between about 100 psia and 450 psia.
 9. The process of claim 1, wherein the first and second portions of the feed gas stream have the same composition.
 10. The process of claim 1, wherein the first portion and second portion of the feed gas streams have a mass ratio in the range of 95:5 to 65:35.
 11. The process of claim 1, wherein the first portion and second portion of the feed gas streams have a mass ratio in the range of 95:5 to 70:30.
 12. The process of claim 1, wherein a portion of compressed mixed refrigerant stream is removed as a supplemental product stream.
 13. The process of claim 1, wherein separating the vapors and liquids in the separator further includes producing a side draw fraction.
 14. The process of claim 13, wherein the overhead vapor stream is enriched in nitrogen and depleted in propane, the bottoms fraction is depleted in nitrogen and enriched in propane, and the side draw fraction has intermediate propane and nitrogen content.
 15. The process of claim 1 further comprising reboiling a portion of the distillation column bottoms in a distillation column reboiler, wherein the energy input to the distillation column reboiler is at least 5% lower than the energy input for a process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, and in which no second portion is formed from the feed gas stream.
 16. The process of claim 1 further comprising reboiling a portion of the distillation column bottoms in a distillation column reboiler, wherein the energy input to the distillation column reboiler is at least 10% lower than the energy input for a process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, and in which no second portion is formed from the feed gas stream.
 17. The process of claim 1, wherein the total compressor duty of the process is at least 5% lower than the compressor duty for a process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, but in which no second portion is formed from the feed gas stream.
 18. The process of claim 1, wherein the total compressor duty of the process is at least 10% lower than the compressor duty for a process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, but in which no second portion is formed from the feed gas stream.
 19. An apparatus for separating natural gas liquids from a feed gas stream, the apparatus comprising: (a) a primary feed gas line configured to deliver a feed gas stream; (b) a heat exchanger operable to provide the heating and cooling necessary for separation of natural gas liquids from a feed gas stream by heat exchange contact between the feed gas stream and one or more process streams thus forming a cooled feed gas stream; (c) a distillation column configured to receive the feed gas stream and to separate the feed gas stream into a column overhead stream comprising a substantial amount of the lighter hydrocarbon components of the feed gas stream and a column bottoms stream comprising a substantial amount of the heavier hydrocarbon components; (d) a first separator configured to receive the distillation column overhead stream and to separate the column overhead stream into an overhead sales gas stream and a bottoms stream comprising a mixed refrigerant configured to provide process cooling in the heat exchanger; (e) a compressor configured to compress the mixed refrigerant stream after the mixed refrigerant stream has provided process cooling in the heat exchanger; and (f) a feed gas bypass line configured to remove a portion of the feed gas stream prior to it being sent to the heat exchanger, wherein the feed gas bypass line is fluidly connected to the distillation column at a point one or more vapor-liquid equilibrium stages below the point at which the cooled feed gas stream from the heat exchanger is fluidly connected, thereby allowing mass transfer exchange between the liquids of the cooled feed gas stream from the heat exchanger and the vapors of the bypass feed gas stream within the column.
 20. The apparatus of claim 19, further including a second separator configured to receive the compressed mixed refrigerant stream and separate the compressed mixed refrigerant into an overhead stream and a bottoms stream that is fed to the distillation column as a reflux stream.
 21. The apparatus of claim 19, further including a splitter configured to provide that the stream entering the feed gas bypass line has the same composition as the portion of the feed gas stream sent to the heat exchanger.
 22. The apparatus of claim 19, wherein the splitter is configured to provide that the bypass line receives about 5 to 35 weight % of the feed gas from the primary feed gas line.
 23. The apparatus of claim 19, further including a splitter configured to remove a portion of the compressed mixed refrigerant as a product stream.
 24. The apparatus of claim 19, wherein the first separator is an absorber.
 25. The apparatus of claim 24, wherein the absorber has a side draw line configured to remove a side draw stream. 